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Exergy and economic analyses of advanced IGCC–CCS and IGFC–CCSpower plants

Nicholas S. Siefert ⇑, Shawn LitsterMechanical Engineering Department, Carnegie Mellon University, United States

h i g h l i g h t s

" Exergy and economic analyses of two advanced fossil fuel power plants configurations." Adv. IGCC–CCS case includes H2 and O2 separation membranes and CO2 capture." Adv. IGFC–CCS case includes catalytic gasifier, pressurized SOFC and CO2 capture." Compared values of IRR and LCOE with conventional power plant configurations." Varied price of NG and CO2 emissions to determine configuration with lowest LCOE.

a r t i c l e i n f o

Article history:Received 13 September 2012Received in revised form 8 January 2013Accepted 2 February 2013

Keywords:Fossil fuelCoal gasificationSolid oxide fuel cellCarbon capture and sequestrationPower plant economics

a b s t r a c t

We present exergy and economic analyses of two advanced fossil fuel power plants configurations: anintegrated gasification combined cycle with advanced H2 and O2 membrane separation including CO2

sequestration (Adv. IGCC–CCS) and an integrated gasification fuel cell cycle with a catalytic gasifierand a pressurized solid oxide fuel cell including CO2 sequestration (Adv. IGFC–CCS). The goal of the exergyanalysis was to evaluate the power generation and the exergy destruction of each of the major compo-nents. We estimated the capital, labor, and fuel costs of these power plants, and then calculated the inter-nal rate of return on investment (IRR) and the levelized cost of electricity (LCOE). In the Adv. IGFC–CCScase, we chose a configuration with anode gas recycle back to the gasifier, and then varied the SOFC pres-sure to find the optimal pressure under this particular configuration. Using a base load generation price ofelectricity of $50/MWh, the IRR of the Adv. IGFC–CCS configuration was 4 ± 3%/yr if the CO2 can be usedfor EOR and 1 ± 3%/yr if the CO2 can only be sequestered in a saline aquifer. The IRR of the Adv. IGCC–CCSconfiguration with H2 and O2 membrane separation was 8 ± 4%/yr if the CO2 can be used for enhanced oilrecovery (EOR) and 3 ± 3%/yr if the CO2 must be sequestered in a saline aquifer. The uncertainty herereflects the uncertainty in capital costs, operation & maintenance (O&M) costs, and CO2 sequestrationcosts.

One goal of the economic analysis was to compare the IRR and LCOE of these configurations with theIRR and LCOE of other fossil fuel power plant configurations. For example, using capital/labor/mainte-nance cost estimates from the literature, we calculated the IRR and LCOE of conventional fossil fuel powerplant configurations, including scenarios with CCS and scenarios with varying costs to emit CO2. We alsopresent results on which power plant configuration yields the lowest value of LCOE as a function of theprice of CO2 emissions and a function of the price of natural gas, holding all other variables constant. AnAdv. IGCC–CCS–EOR configuration yields the lowest value of LCOE when the price of natural gas and theprice of CO2 emissions are above the line between ($5/GJ, $10/tCO2) and ($2.5/GJ, $50/tCO2).

Published by Elsevier Ltd.

1. Introduction

Pulverized coal combustion (PCC) power plants generate be-tween 40% and 50% of the total supply of electricity in the United

States [1]. However, this percentage is likely to decrease in the fu-ture because of the currently low price of natural gas as well as therecent proposed regulations on the emission of greenhouse gasesby the Environmental Protection Agency [2]. While the future forbuilding new PCC power plants looks bleak, the future may notbe as bleak for building advanced integrated gasification combinedcycle with carbon capture and sequestration (IGCC–CCS) and

0306-2619/$ - see front matter Published by Elsevier Ltd.http://dx.doi.org/10.1016/j.apenergy.2013.02.006

⇑ Corresponding author.E-mail address: [email protected] (N.S. Siefert).

Applied Energy 107 (2013) 315–328

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Applied Energy

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advanced integrated gasification fuel cell with carbon capture andsequestration (IGFC–CCS) power plants that operate off of mixturesof coal, municipal solid waste, or petroleum coke.

Herzog [3] published in 1999 one of the first reports detailingthe economic costs of carbon dioxide capture at coal power plants.Since then, there have been a number of economic analyses of ad-vance fossil power plants with and without carbon dioxide capture,such as Johnson and Keith [4], Rubin et al. [5,6], Davison [7],Patino-Echeverri et al. [8], Kunze and Spliethoff [9], Hu et al.[10], Hammond et al. [11], Fischbeck et al. [12], Viebahn et al.[13], Pettinau et al. [14] and Melchior and Madlener [15]. Johnsonet al. [4] and Fischbeck et al. [12] have analyzed the effect of bothnatural gas prices and the price of carbon dioxide emissions on theeconomic viability of the various fossil fuel power plant configura-tions. There have also been numerous studies on the economic via-bility of various fossil fuel power plants configuration conductedby the National Energy Technology Laboratory (NETL) and theEnergy Information Administration (EIA). Some recent NETL stud-ies on the economics of various advanced coal power plants withCCS, such as by Gerdes et al. [16,17] and Grol and Wimer [18], in-clude capital cost estimates of advanced IGFC–CCS power plantconfigurations. The Adv. IGCC–CCS and Adv. IGFC–CCS configura-tions analyzed here are similar to the configurations modeled byGerdes et al. [16,17] and Li et al. [19], with the main differencebeing that Li et al. [19] included the sale of both electricity andhydrogen.

Herein, we present our exergy and economic analyses of twodifferent advanced gasification based power plants where themain product was electricity. The first configuration analyzedis similar to a conventional IGCC–CCS configuration [16,20], butwith a few noticeable changes: (1) ion transport membranes(ITMs) for O2 separation rather than cryogenic air separation;(2) warm gas sulfur removal with alkali hydroxide rather thanlow temperature removal using physical solvents; and (3) palla-dium membranes for H2 separation from the syngas rather thanlow temperature removal of CO2 from the syngas using physicalsolvents. The second configuration analyzed is an advancedIGFC–CCS configuration in which a catalytic coal gasifier is cou-pled with a pressurized hybrid solid oxide fuel cell featuring acompressor and turbine pair that can operate at pressures be-tween 0.2 and 0.8 MPa. For both configurations, we calculatedthe exergy efficiency, the internal rate of return on investment(IRR), and the levelized cost of electricity (LCOE). We then com-pared the values of IRR and LCOE to other fossil fuel powerplants. The goal of the exergy analysis in this paper was not tooptimize the power plant for exergy efficiency, but rather wasto determine where in the power plant exergy is lost due to irre-versible processes.

In addition to conducting exergy and economic analyses ofthese particular configurations, we compare the cost of electricityof these two advanced power plants with power plant configura-tions designed to meet potential Environmental ProtectionAgency (EPA) regulations that limit greenhouse gas emissions fromnew large-scale power plants to less than 0.45 kg (1 lb) of CO2

per kW h of electricity generated. To meet these proposedrequirements, a PCC or an IGCC power plant needs to captureand sequester roughly 50% of the carbon dioxide produced atthe power plant. We label these power plant configurationsPCC–50%CCS and IGCC–50%CCS. In this paper, we compare theIRR and LCOE calculated in this report to the IRR and LCOE ofother power plant configurations analyzed by Rubin et al. [5] andGerdes et al. [16,17]. In addition, we analyze at what price ofnatural gas and at what price of CO2 emissions can advanced coalbased power plants with CCS compete economically with naturalgas combined cycle power plants (NGCCs).

2. System and exergy analysis

Exergy and thermo-economic analyses date back to at least the1960s [21], and have been reviewed by El-Sayed [22]. Exergy is nota standard thermodynamic variable because the exergy of a systemis given with respect to the temperature, pressure and compositionof the Earth’s atmosphere, which is not the same at different timesor locations. Our assumed standard state is the following: 25 �C,0.10 MPa, and a molar gas composition of 78% N2, 20% O2, 2%H2O, and 0.04% CO2. The exergy of a system is the maximum usefulwork that can be generated during a process that brings the systeminto thermal, chemical and mechanical equilibrium with the sys-tem’s environment. Exergy can exist in multiple forms, such asthermo-mechanical exergy and chemical exergy; there are alsocorresponding forms of exergy for kinetic, potential, and electricalenergy. If ignoring the potential and kinetic exergy of the flow, themolar exergy of a substance flowing into or out of a control volumecan be defined as:

e ¼ ðh� h0Þ � T0ðs� s0Þ þ Rixiðli � li0Þ ð1Þ

where h is the molar enthalpy, T is the temperature, s is the molarentropy, xi is the mole fraction of species i, and li is the chemicalpotential of species i at standard temperature and pressure. Termswithout the naught symbol are for the system, and terms withthe naught symbol are for the environment. The exergy of a sub-stance cannot be negative, and it is only equal to zero when the sub-stance is in thermal, chemical and mechanical equilibrium with itsenvironment. By combining the first and second laws of thermody-namics for open, steady-state processes, one obtains the exergy bal-ance equation:

_Wuseful ¼ Ri _nieþ RkTk � T0

Tk

� �_Qk � T0 � _rirr ð2Þ

where _Wuseful is the amount of useful mechanical and electrical workgenerated from the system, _ni is the molar flow rate into or out ofthe control volume of component i, Tk is the temperature at whichheat _Qk flows out of or into the control volume, T0 is the referencetemperature of the environment, and _rirr is the rate of entropy gen-eration inside of the control volume due to irreversible processes. Itis these irreversible processes, such as the flow of particles acrossgradients in temperature, pressure or chemical composition, whichcause exergy destruction. When exergy is destroyed, there is a lossin the amount of useful work available that can be generated fromthe original exergy available to the system. The amount of exergydestruction is given by the Gouy–Stodola theorem [23]:

Udes ¼ T0 � rirr ð3Þ

where Udes is the exergy destruction and rirr is the amount of entro-py generated by irreversible processes. Some of the value in con-ducting an exergy or second law analysis, in addition to a firstlaw analysis, are the following: (1) identification of how muchpotentially useful work is destroyed within processes within theplant, (2) ensuring that none of the individual idealized processesinside of the plant violate the second law of thermodynamics, and(3) ensuring that the total exergy entering the plant is equal exactlyto the amount of exergy leaving the plant plus the exergy destruc-tion inside of the plant. While a standard exergy analysis is animportant check to understand where improvements in the plantmight be possible, a standard exergy analysis does not include cal-culations of the cost and/or exergy destruction associated withbuilding and maintaining the power plant. Therefore, while an exer-gy analysis is useful in estimating the cost to fuel the power plant,an exergy analysis is not a substitute for a full economic analysis ofa particular power plant configuration.

316 N.S. Siefert, S. Litster / Applied Energy 107 (2013) 315–328


2.1. Adv. IGCC with H2 and O2 separation membranes

2.1.1. Description of process flow diagramWe conducted an exergy analysis of an integrated gasification

coal power plant with advanced hydrogen and oxygen separationmembranes. Mass and energy balances for each of the individualreactors in the system were conducted using HSC Chemistry 6.0(Outotec, Espoo, Finland), which also calculated the chemical equi-librium composition given input flows by minimizing the Gibbsfree energy. Mass, energy and exergy balances were conductedusing Cantera v1.7, which is open source software by D.G. Good-win. Fig. 1 shows the process flow diagram for all the flows andequipment modeled for this configuration. The following sub-sec-tions provide a description of the technologies used in the processand highlight some important technological details.

2.1.1.1. Gasification. We modeled a GE entrained flow gasifier usingHSC Chemistry. Coal is crushed and then mixed with water beforeentering a slurry pump to pressurize the slurry to 4.2 MPa. The tem-perature at the exit of the gasifier is 1200 �C, but after the syngas cool-er, the temperature is 860 �C. We used a H2O-to-carbon ratio and anO2-to-carbon ratio in the gasifier of 0.46:1 for both H2O and O2, whichgenerates a syngas composition of 30% H2, 48% CO, 15% H2O, and 7%CO2 using Gibbs free energy minimization in HSC Chemistry.

2.1.1.2. Syngas quench, sulfur removal and water gas shift. After exit-ing the syngas cooler, the syngas is quenched with liquid watercontaining sodium hydroxide. The sodium hydroxide is used tocapture acid gases, such as hydrogen sulfide and carbonyl sulfide.We assumed the presence of 0.5% molar composition of H2S plusCOS in the syngas for the economic model. We assume that thehydrogen sulfide and carbonyl sulfide is converted into elemental

sulfur via either the Thiopaq (Paqell, Balk, The Netherlands) andLO-CAT (Merichem, Houston, TX) processes.

After exiting the quench reactor, the syngas is saturated withwater at a temperature of 250 �C, and then is sent to a bed of acti-vated carbon to remove mercury and to remove any further H2S inthe gas stream. The syngas then enters a water gas shift (WGS)reactor at a constant temperature of 250 �C and a constant pressureof 4.2 MPa. The syngas composition exiting the WGS reactor is 54%H2, 3% CO, 7% H2O, and 37% CO2, using Gibbs free energy minimi-zation constrained to not form methane in the reactor. The thermalenergy from the WGS reactor is removed via heat transfer withsteam in the Rankine cycle.

2.1.1.3. Hydrogen separation. There are various methods of separat-ing hydrogen from syngas streams using inorganic membranes[24]. This technology is still in the early stages of commercialdevelopment. Typically, palladium is used because of its high per-meability for hydrogen diffusion through the solid. The palladiumis normally doped with other metals, such as copper, in order to re-duce the cost of the membrane and to increase the tolerance of thealloy to hydrogen sulfide poisoning [25,26]. The flux through themembrane was estimated using data from prior research on palla-dium alloy membranes [27] chosen at the temperature after thesyngas compressor of 726 K. The hydrogen pressure on the pureside of the membrane is assumed to be to be 0.5 MPa, yielding anormalized flux of hydrogen through the membranes of roughly0.02 mol m�2 s�1.

2.1.1.4. Carbon dioxide capture. The hydrogen-depleted streamfrom the palladium membrane reactor is oxy-combusted with justenough oxygen to convert all carbon monoxide into carbon dioxideand all remaining hydrogen into water vapor. After catalytic

Fig. 1. Process flow diagram of the IGCC–CCS process modeled. Left: coal gasification and quench, Middle: water gas shift and H2 separation, Bottom: Brayton cycle and O2

separation, and Top Right: Rankine cycle.

N.S. Siefert, S. Litster / Applied Energy 107 (2013) 315–328 317


oxy-combustion so that there is no oxygen remaining in the gasstream, the gases are cooled such that they leave the heat exchan-ger as liquid water and supercritical carbon dioxide. The carbondioxide is separated and pumped to an existing CO2 pipeline thatis assumed to be located 50 km from the power plant.

2.1.1.5. ITM oxygen separation. In order to increase the system effi-ciency and decrease the capital cost compared to an IGCC–CCS con-figuration with cryogenic air separation, Air Products andChemicals, Inc. (Allentown, PA) is currently developing ion trans-port membranes (ITMs) to provide the oxygen for the gasifier[28]. In this Adv. IGCC–CCS configuration, oxygen is separated foruse throughout the process by sending the hot, compressed gasfrom the exit of a Brayton cycle compressor to the ITM separationreactor. We assume that the mixed ionic–electronic ceramics of theITM are 100% selective in separating oxygen from air. This processrequires temperatures between 800 �C and 900 �C, and a sizeablepressure difference [29], yielding a normalized flux of oxygen ofroughly 0.03 mol m�2 s�1. The pressure on the air side of the mem-brane is the same as the pressure of the Brayton cycle turbine(1.7 MPa). We assume that the pressure of oxygen on the pure oxy-gen side of the membrane is 0.1 MPa, and therefore, the oxygenmust be cooled and compressed to 4.2 MPa before entering thegasifier and compressed to 15.0 MPa before entering the oxy-combustor.

2.1.1.6. Brayton cycle. The Brayton cycle turbine modeled here isthe GE 9001FA model with a power output of 255.6 MW, a pres-sure ratio of 17.0 and a heat rate of 9757 kJ per kW h. The model

assumes the plant operates two of these turbines for a total Bray-ton cycle power output of 511 MW. Using this publically availabledata on the GE 9001FA, we calculated that the isentropic efficiencyof the compressor and turbine of the Brayton cycle was 86%. In thecombustor, hydrogen reacts with air from the main compressorand with the depleted air from the ITM oxygen separation process.The pure hydrogen stream leaving the palladium membranes iscompressed from 0.5 MPa to the pressure of the Brayton turbinecombustor (1.7 MPa). The main system compressor providesapproximately 235% excess air. This limits the adiabatic flame tem-perature of the combustor to 1430 �C, the specified firing temper-ature for the turbine. The combusted gases pass through the vanesof the turbine and are then sent to the steam generator for the Ran-kine cycle.

2.1.1.7. Rankine cycle. The main Rankine steam generator utilizesthe thermal energy from the Brayton cycle exhaust, the thermalenergy from the carbon capture oxy-combustion exhaust, the ther-mal energy from the water gas shift reactor, and the thermal en-ergy obtained by cooling the pure oxygen stream before thegasifier. This thermal energy is used to generate the steam neces-sary to drive the high-pressure (HP) steam turbine. The HP turbineexhaust is sent through a portion of the steam generator to providethe reheat necessary for the intermediate pressure (IP) turbine. Theexhaust of this turbine is sent directly through the low-pressure(LP) turbine. Here, we consider a GE 207FA steam turbine with aHP stage at 13 MPa/565 �C, an IP stage at 3 MPa/565 �C, and anLP stage at 0.5 MPa with no reheat. We assumed an isentropic effi-ciency of 90% when modeling the Rankine cycle. The three stageshave a total output of 459 MW.

2.1.2. Technical performanceAs seen in Table 1, the total power output was calculated to be

808 MW. This accounts for electricity generated by the two gas tur-bines as well as the Rankine cycle steam turbine. It also includesplant electricity requirements for all major system loads presentin the process diagram in Fig. 1. Using the assumptions listedabove, the power plant operates at a HHV thermal efficiency andexergetic efficiency of around 43%. For comparison, NETL’s analysisof an IGCC–CCS power plant with ITM, H2 Sep & G9FA is quite

Table 1System output variables for the Adv. IGCC–CCS configuration model in this report.

Brayton cycle power output 511 MWRankine cycle power output 459 MWTotal power output 808 MWCoal input rate 5230 ton/dayCO2 produced 14,640 ton/dayCO2 capture rate 14,640 ton/dayThermal efficiency (HHV) 43.4%Exergetic efficiency 42.8%

Catalytic Gasifier3.0 MPa, 700oC

Coal

Solid Oxide Fuel Cell

0.2 -0.8 MPa, 800 -900 oC

Air

Exhaust Air CH4, H2, CO, H2O, CO2

CO2, H2O, H2, CO

Cathode

Anode

CO2 (15 MPa)H2O, CO2

K2CO3 ,KSiAlO4 & K2S

KOH

Insoluble Catalyst Regeneration via

ElectrodialysisAsh: SiO2, Al2O3

CO2 Capture with MgO/CaO(HX#1)

HX#2

CO2 release from (Mg,Ca)CO3

(HX#2)

HX#2

HX#1

HX#3

18

2221

23

24

1

3

5

16

H2O

6

7

8

9

10

Soluble Catalyst

Regeneration

K2CO3K2S

HX#3

Coal & Catalyst Mixing

KSiAlO4

2

4

11

12

13

14

15

17

1920

HX#3

Gas

Liquid

Solid

Fig. 2. Process flow diagram for catalytic steam gasification integrated with a SOFC.



similar to the flow diagram studied here [16], in which they mea-sured a HHV efficiency of 40% as seen in Fig. 3, which will be dis-cussed further in section 2.3.

2.1.3. Exergy analysisIn this section, we analyze each of the main locations of exergy

destruction (irreversible entropy generation) within the powerplant. Table 2 shows a list of sources for the exergy destruction,as well as the amount of exergy leaving the power plant, normal-ized by the inlet exergy flow of the incoming coal. We calculatedthat 42.8% of the inlet exergy will leave the system as electricity,6.2% will leave as exergy in the compressed carbon dioxide, andonly 0.9% will leave as the exhaust’s thermal exergy. The largestsources of exergy destruction are the coal gasifier/syngas coolersection (19.6%), the combustor of the Brayton cycle (12.4%), thesteam turbine (4.2%), the gas turbine (2.9%), WGS reactor withassociated heat exchanger (HX) (2.9%), Post-Brayton Cycle HX(2.1%), the quench system (2.0%), the H2 membrane separation pro-cess (1.4%), the gas compressor (0.9%), the ITM O2 separation mem-brane reactor (0.2%), and finally the Rankine Cycle Pump (0.1%).The exergy destruction in the gasifier was significant because thereis an inherent mismatch between the temperature of gasification(1200 �C) and the temperature at which the heat is transferred tothe steam in the Rankine cycle (<600 �C). There was also significantexergy destruction in the combustor because, even though thetemperature of the combustor is quite high, this is the locationwhere most of the fuel is oxidized.

2.2. Adv. IGFC with catalytic gasifier and pressurized SOFC

2.2.1. Description of process flow diagramWe conducted an exergy analysis of a power plant design in

which a catalytic coal gasifier produces a methane rich syngas. Car-bon dioxide is captured from the syngas before the syngas is sentto a SOFC. In this design, the anode tail gas from the SOFC is recy-cled back to the catalytic gasifier. Fig. 2 shows the process flow dia-gram for the major components of this system. We used HSCChemistry for both a first law balance and a material balance forthe catalytic gasifier, the CO2 capture/release reactors, and theSOFC. The gas compositions throughout the loop of Steps 4–9 arelisted in Table 3 when the single pass utilization of the SOFC was70%. Since this is a process with a recycle loop and with multiplespecies capable of being oxidized in the fuel cell, ‘single pass fuelutilization’ is defined to be equal to the amount of oxygen thatcrosses from the cathode to the anode divided by the amount ofoxygen that would cross from the cathode to the anode if all ofthe H2, CO, and CH4 entering the anode were completely oxidizedto H2O and CO2.

2.2.1.1. Catalytic coal gasifier. The catalytic gasifier modeled here isbased off of the Exxon single-stage, fluidized bed catalytic gasifier[30,31], in which low rank coals are slurry mixed with roughly20 wt.% potassium hydroxide and carbonate. A slurry of coal, cata-lyst and water is pumped to the pressure of the gasifier (3 MPa)and then dried using exhaust air from the low pressure Brayton cy-cle. This catalytic gasifier operates adiabatically at a temperature of700 �C. The dried coal and catalyst enter the catalytic gasifier alongwith the anode gas recycle from the SOFC. The gasifier is operatedadiabatically. The molar methane composition of the syngas fromthe catalytic gasifier is roughly 20% on a dry basis [30,31], becauseof non-equilibrium effects such as methane volatilization from coal[32], but these effects could not be modeled in a chemical simula-tor that minimizes the Gibbs free energy. In addition to catalyzingthe steam–coal gasification reaction, the alkali catalyst also cancapture acid gases, such as hydrogen sulfide [33], which simplifiesthe syngas cleanup steps before the SOFC.

2.2.1.2. Catalyst regeneration. The catalyst, ash and unconsumedcarbon exit the gasifier and are quenched with water. We assumethat the residence time in the gasifier is such that there is only 1%unconsumed carbon compared with the initial energy content ofthe coal. After the catalyst, ash and unconsumed carbon arequenched with water, and soluble species will enter the aqueousphase, such as potassium carbonate, potassium sulfide and somepotassium alumina-silicate species. Yeboah et al. [34] and Shethet al. [35] have studied the effect of different catalysts on the gas-ification rate, and determined that when the potassium is bondedto weak anions (such as OH�, S2� and CO2�

3 ), the kinetics rates ofgasification were higher than when bonded to strong anions (suchas Cl�) that are stable in gasification environments. Since potas-sium carbonate and potassium sulfide are both water soluble andactive catalysts, the water soluble catalysts can be recovered at thisstage and mixed with fresh coal.

When alkali cations react with alumina-silicates in the coal ash,their ability to catalyze steam–coal reactions significantly de-creases [36]. Since alkali cations thermodynamically prefer beingchemically bonded to alumina-silicates than being bonded withcarbonate anions or sulfide anions, the alkali cations must berecovered from alkali alumina-silicates in order to maintain thecatalytic capability of the alkali ions.

H2OðlÞ þ K2AlxSiyOzðsÞ ¼ AlxSiyOzðsÞ þ 2KOHðaÞDG300K � þ12 kJ=mol ð4ÞFig. 3. First law system efficiency for various power plant configurations with

greater than 90% CO2 capture and compression to 15 MPa.

Table 2Exergy destruction and exergy output as a percentage of the inlet exergy for an Adv.IGCC–CCS configuration with H2 and O2 membrane separation.

Net power exiting 42.8%Gasifier RC HX 19.6%Brayton cycle combustor 12.4%SC CO2 exiting 6.2%Steam turbine 4.2%WGS reactor and RC HX 2.9%Brayton cycle turbine 2.9%Post-brayton cycle HX 2.1%Syngas quench 2.0%Rankine cycle condenser 1.4%H2 membrane + H2 compressor 1.4%Brayton cycle compressor 0.9%Exhaust air 0.9%O2 membrane 0.2%Rankine Cycle Pump 0.1%Sum 100.0%



To determine the amount of electricity required to generate alkalihydroxides from the alumina-silicates, we first calculated theamount of alkali species that react with alumina-silicates in the gas-ifier based on the ash content of the coal. For this configuration, weassume that the coal is a low ash, low sulfur coal from Power RiverBasin, WY with a weight fraction of 0.5% kg kg�1 sulfur and 5%kg kg�1 alumina-silicate. We then estimate the amount of electricityconsumption by multiplying the molar flow of alkali alumina-sili-cates times theDG of Eq. (4), and then divide this quantity by an elec-trodialysis electrical efficiency of 40%. We chose the value of 40%because it falls within the range of the electrodialysis efficienciesmeasured experimentally by Nagasawa et al. [37].

As stated earlier, the actual consumption of electricity in theelectrodialysis unit will depend on the ash content of the coal. Aswill be shown in the exergy analysis section, catalyst regenerationusing bipolar membrane electrodialysis consumes roughly 1% ofthe gross electricity generated at the power plant if a low ash coalis used as fuel and if the weight of the alkali carbonate catalyst is0.2 kg for every 0.8 kg of coal used.

2.2.1.3. Carbon dioxide capture. There are various methods ofremoving carbon dioxide from coal gasification syngas [38]. Com-mercially available physical solvents, such as Selexol (UOP LLC)or Rectisol (Linde AG and Lurgi AG), require lowering the temper-ature to below the dew point of the syngas. Instead, we model achemical capture process that leaves the temperature of the syngasclose to the inlet temperature of the solid oxide fuel cell. After leav-ing the catalytic gasifier, the methane rich syngas goes through anexpander to drop the pressure of the gas to the pressure of theSOFC. Then the gas goes to a reactor filled with magnesium and cal-cium oxide (MgO, CaO) in order to capture CO2 as well as anyremaining H2S and COS in the gas stream [39]. Carbon dioxide cap-ture occurs at a temperature of 750 �C or less, depending on thepressure after the expander. The CO2 is regenerated from the dolo-mite (MgCO3, CaCO3) using hot exhaust gases from the SOFC at atemperature of 1000 �C and a pressure of 0.1 MPa. After this, thecarbon dioxide is cooled, compressed, cooled, dried, and then com-pressed to a pressure of 15 MPa for subsequent injection into a car-bon dioxide pipeline.

2.2.1.4. SOFC. The syngas leaving the CO2 capture reactor then en-ters the anode of the SOFC. The SOFC is the main source of electric-ity generation from this power plant configuration. The SOFC ismodeled using V–i curves at various SOFC temperatures and pres-sures using publically-available data from Rolls Royce Fuel CellSystems [40]. The equation used to model the fuel cell voltagewas the following:

V ¼�gf ;H2OðgÞðT;1 atmÞ

2Fþ RT

2Fln

panodeH2ðgÞ � pcathode

O2ðgÞ

� �1=2

panodeH2OðgÞ

264

375

� i � ASR� RTðazÞ � F ln

i

i0o

þ 1

!ð5Þ

The voltage, V, between the anode and the cathode is equal tothe open circuit voltage minus the Ohmic overpotential minusthe electrode overpotential, where i is the operating current den-sity in [A cm�2] and (az) is the transfer coefficient, which we as-sume to be equal to a value of 2 in order to estimate theelectrode exchange current density, i0

o . The Gibbs free energy of for-mation of water from H2 and O2 as function of temperature at1 atm, gf,H2O(g) (T, 1 atm), was determined using HSC Chemistry.The pressures are the average pressure along the length of the fuelcell and the units of pressure in the equation above are atmo-spheres. The values of the Ohmic area specific resistance, ASR,and the electrode exchange current density, i0

o , are both functionsof temperature, and the electrode exchange current density is alsoa function of pressure. We determined the values of ASR and i0

o byfitting publically available data from Rolls Royce Fuel Cell Systems[40]. Since the Rolls Royce data was given as both a function oftemperature and pressure, we were able to calculate the followingvalues of ASR and i0

o .

ASR ½X cm2� ¼ 0:12þ 0:18 � e6500� 1T�

11123ð Þ: ð6Þ

i0o ½A � cm2� ¼ 0:01 � p½atm� � e6500� 1

T�1

1123ð Þ ð7Þ

where T is the temperature in Kelvin. Using this empirically fit data,we were able to calculate the voltage of the fuel cell as a function oftemperature, pressure and gas composition.

After leaving the fuel cell, as seen in Fig. 2, most of the anode tailgas goes directly to HX2; however, a small portion of the tail gas ismixed with the depleted air exiting the cathode. This is effectivelya bleed stream in order to prevent the build of inert gas species,such as N2 in originally in the coal as nitrogen species, and to pre-vent the buildup of water vapor.

Gaseous fuels like methane, ammonia and carbon monoxideare internally reformed or shifted in the anode channels of theSOFC to yield the hydrogen that reacts with oxygen ions onthe anode [41,42]. There have been a number of research groupsthat have demonstrated experimentally the capability of dopedNi–YSZ anodes to reform methane and higher hydrocarbons[43–49]. For example, Shiratori et al. [46] experimentally demon-strated operation of a SOFC for 50 h with direct biogas using aNi–ScSZ cermet as the anode material. While most anodes com-posed of pure Ni–YSZ are not tolerant to high levels of H2S or tohydrocarbons, Yang et al. [48] showed that Ni–YSZ anodes dopedwith barium and cerium are more tolerant to both hydrogen sul-fide and propane.

2.2.1.5. Brayton cycle. Depending on the fuel cell pressure, therecan be significant net power generation from the combined aircompressor and exhaust turbine. In this configuration, the air isfirst compressed and then sent to the CO2 capture reactor to pro-vide the cooling required to maintain the temperature for CO2 cap-ture below 750 �C. The air then enters the cathode of the fuel cell.After the cathode, the air combines with the portion of the anodetail gas that is not recycled back to the catalytic gasifier. This ex-haust gas is combusted, raising the temperature of the exhaustto the point at which it can be heat exchanged with the magnesiumand calcium carbonate exiting the CO2 capture reactor. After HX#2,the exhaust air passes through an exhaust turbine. The isentropicefficiency of all compressors and turbines in this system was as-sumed to be 80%.

2.2.2. Exergy analysisWe now focus on the exergy analysis on this configuration that

integrates a catalytic gasifier with a pressurized fuel cell operatingon a methane-rich syngas. We calculated an exergetic efficiency of58.3% for the operating conditions listed in Table 4: catalytic

Table 3Syngas composition throughout the loop that integrates the catalytic gasifier with theSOFC, as calculated by HSC Chemistry using Gibbs free energy minimization, whenthe single-pass fuel utilization was 70%.

Gasifier exit(4) (%)

Post CO2

capture (6) (%)Post SOFCanode (7) (%)

H2 32 58 29CO 13 2 6CO2 13 1 6H2O 34 15 59CH4 8 14 0



gasifier pressure was 3.0 MPa; SOFC pressure was 0.5 MPa; air stoi-chiometric ratio was 2.0; SOFC current density was 0.5 A cm�2;SOFC Voltage was 0.70 V; and SOFC single pass fuel utilizationwas 70%. In Table 4, we list where power is either generated orconsumed as well as where exergy is destroyed in the power plantdue to irreversible processes. Our calculations in Table 4, as well asthose by Li et al. [50], show that the system efficiency for powerplants with catalytic gasification with anode recycle is near 60%.

As seen in Table 4, the largest source of exergy destruction wasthe CO2 capture using a combination of calcium and magnesiumoxide and the associated heat exchangers to cool or heat the solidmaterials. The second largest source of exergy destruction was theSOFC. The exergy destruction inside the SOFC is due to irreversibleprocesses within the fuel cells, principally Ohmic and cathode acti-vation losses. Using the Gouy–Stodola Theorem as presented in Eq.(3), the exergy destruction of a methane-fueled SOFC at constanttemperature is roughly equal to the overvoltage, g, times the currenttimes the temperature of the environment (298 K), To, divided by thetemperature of the SOFC, TSOFC, and divided by the exergy of the fuelinto the power plant normalized by the fuel’s reduction charge [33].

SOFC exergy destruction½%� ffi g1 V� T0

TSOFCð8Þ

Depending on the fuel, the exergy divided by the reductioncharge, i.e. the number of electrons generated if the fuel is fullyoxidized on an electrode, is typically between 1.0 and 1.3 V. Giventhe amount of electricity generated in the SOFC, the SOFC is not amajor source of exergy destruction because the temperature of theSOFC is nearly four times larger than the temperature of theenvironment.

The third largest source of exergy loss is the 150 bar CO2 leavingthe power plant. This high pressure is required to overcome fric-tion in the pipeline and to overcome the pressure of a typical geo-logic reservoir. Here, we count the exergy required to sequester theCO2 as exergy destruction because the pressurized carbon dioxideis not being used to generate electricity at the power plant, eventhough this mechanical form of exergy is used for useful purposesin the EOR case. There was only minor exergy destruction inside ofthe catalytic gasifier because there is no oxygen consumption in-side of the catalytic gasifier. The exergy destruction of the turbinesand compressors were each 3% or less.

2.3. Comparison with other researchers

It has been shown by Gerdes et al. [16,17], Grol and Wimer [18],Shelton et al. [51], and Li et al. [50] that coal-based power plants

using catalytic gasifiers and pressurized fuel cells can achievesystem efficiencies of �60% while sequestering >90% of the carbondioxide generated at the power plant. Though, there are some nota-ble differences in the approaches in each of the studies listed above.A few key differences between the Adv. IGFC system analyzed hereand some of the Adv. IGFC systems analyzed by previous researchare: (a) anode tail gas recycle back to the gasifier; (b) CO2 capturebefore the SOFC; (c) no steam turbine; and (d) intermediate SOFCpressure. The SOFC pressure in other system analyses has eitherbeen 0.1 MPa or greater than 1 MPa. Here, we analyzed casesbetween 0.2 MPa and 0.8 MPa. We chose to capture CO2 beforethe SOFC because this CO2 capture step will reduce the chance ofcarbon build-up on the anode electrode and also remove the major-ity of any remaining sulfur species in the syngas. Since IGFC systemdesigns vary significantly between research groups, the optimalconfiguration depends on the exact constraints and costs of fuel cellsystems. Also, since many of the main pieces of equipment in anIGFC power plants are not commercial-off-the-shelf technologyand since their performance and cost are still evolving, there is noway to definitely prove that there is an optimal configuration.

In addition to comparing with similarly designed advancedpower plants, we also compare the system efficiency of the powerplants modeled here with various conventional power plant con-figurations containing greater than 90% carbon capture and com-pression of the CO2 to 15 MPa. Fig. 3 shows the first law systemefficiency (Net Work vs. Higher Heating Value, HHV) for a widerange of coal-fired, base-load power plants. The cases in blue arefrom NETL’s analysis of various coal-based power plants [16]; thecases in red are from Li et al. [50]; and the cases in green are forthe two configurations analyzed in this section.

As seen in Fig. 3, the system efficiency of a conventional pulver-ized coal combustion (PCC) power plant with post-combustion car-bon capture is around 27%. The system efficiency of an IGCC powerplant with pre-combustion carbon capture is between 32% and43%, depending on (a) the method of oxygen separation from air,(b) the method of separation of CO2 from H2, and (c) the tempera-ture of the sulfur removal process. The system efficiency of an IGFCpower plant is between 42% and 58%, depending greatly on (a) thetype of gasifier, (b) the operating voltage, (c) the pressure of theSOFC, and (d) whether there is anode recycle back to the gasifier.One clear trend in the models is that the system efficiency in-creases when the carbon dioxide is captured at elevated pressurerather than at atmospheric pressure. Another trend is that systemefficiency increases for similar configurations when a catalytic gas-ifier replaces a conventional entrained flow gasifier. The questionwe address in the next section of this report is whether the config-urations with higher system efficiency are cost effective comparedwith more traditional PCC-CCS and IGCC–CCS configurations.

3. Economic analysis

3.1. Methodology for economic analysis

While knowing the first or second law efficiency of a powerplant is useful in estimating the costs of fueling a power plant,the first or second law efficiency cannot predict the economic via-bility of the power plant. A detailed knowledge of the capital, fueland labor costs are required in order to calculate a figure of meritwith which to compare the power plant with other investmentopportunities. One should not optimize a power plant to obtainmaximum system efficiency because the fuel-to-electricity systemefficiency does not account for either the cost or the irreversiblegeneration of entropy associated with building, fueling, maintain-ing, and deconstructing the power plant. Therefore, in addition tothe exergy analyses presented earlier, we have also conducted

Table 4Exergy balance for a SOFC operating on syngas from a catalytic gasifier and anode tailgas recycle back to the catalytic gasifier. Pressure of the SOFC was 0.5 MPa; airstoichiometric ratio was 2.0; SOFC current density was 0.5 A cm�2; SOFC voltage was0.7 V; and SOFC temperature was 1123 K (850 �C).

Process step Power/inletexergy (%)

Exergy destruction or loss/inlet exergy (%)

Air compressor (18–19) �12.9 0.8Air turbine (22–23) +13.8 1.1Exhaust air (24) +0.8 2.3CO2 compressor (16–17) �1.4 0.4Exhaust CO2 (17) – 5.5CO2 capture HX’s (5, 6, 7, 8, 19,

20, 21, 22)– 12.9

Catalyst regeneration withelectrodialysis (12–14)

�0.8 0.4

Catalytic gasifier (3, 4, 9, 10) – 3.5Syngas expander (4–5) +10.5 0.7Syngas compressor (8–9) �13.9 3.2SOFC (6–7, 20–21) +62.3 10.8Sum 58.3 41.7



economic feasibility analyses for these two power plant configura-tions. The feasibility analyses are Class 4 capital cost estimates asdefined by the Association for the Advancement of Cost Engineer-ing International (AACE) [52]. The expected accuracy of a Class 4capital cost estimate is �15% to �30% on the low side and +20%to +50% on the high side. This means that there will be significantuncertainty in the actual capital cost of the configurations analyzedin this study, and therefore, the capital costs detailed belowshould be assumed to have the level of uncertain on the order of+50%/�30%.

The goal of this capital cost estimate, as well as the LCOE/IRRanalysis, is to evaluate the economic viability of future technolo-gies that have been modeled here, such as H2 and O2 separationmembranes, catalytic gasifiers, and pressurized SOFCs. Costs forcapital, operation and maintenance were determined through costestimations from the Integrated Environmental Control Model(IECM) [5] for commercial technology and from Gerdes et al.[16,17] for pre-commercial technologies, such as O2 and H2 separa-tion membranes. It should be noted that the actual upfront capitalcost ($/kW) of PCC and IGCC technologies in recent years has beenup to twice the overnight capital cost (2007USD/kW) listed in thepapers from which this report has derived its capital cost esti-mates, and may be due what is called the Averch–Johnson effect[53]. Another reason for this difference may be due to the fact thatcapital cost values in the literature assume equipment productionrates higher than actual production rates. While there is largeuncertainty in any large-scale power plant economic analysis, thegoal here is simply to help determine whether certain technologiesjustify further research and development, and should not be con-strued as support for or against any of companies or technologiesdiscussed above.

One potentially useful economic figure of merit is the levelizedcost of electricity (LCOE). The LCOE includes all of the variables re-lated to building, fueling, operating and decommissioning thepower plant, and in addition, the LCOE includes the interest rateon the capital loans. A simplified equation for calculating the LCOEis given below:

LCOE ¼M þ F þ P þ C � rð1þrÞn

ð1þrÞn�1

h i� ð1þ rÞt þ D � r

ð1þrÞn�1

h iE

ð9Þ

where M is the yearly operations and maintenance expenditures; Fis the yearly fuel expenditures; P is the yearly pollution creditexpenditures; C is the upfront capital expenditures; t is the con-struction time, weighted to account for how funds are spent duringstart-up; D is the decommissioning investment expenditures; r isthe discount/interest rate on the capital loans; n is the number ofyears the system is operational; and E is the net yearly electricitygeneration. In this report, we will compare the values of LCOE wecalculate with the value of LCOE for other base-load power plants,as calculated by previous studies referenced above. For the LCOEanalyses in this report, we use an inflation-adjusted discount rateof 7%/yr. This value was chosen because it is the suggested valuefor an inflation-adjust discount rate is 7%/yr, according to the US Of-fice of Management and Budget (OMB) [54]. Although, it should benoted that this OMB guideline has not been updated since 1992 andwe recognize that the choice of discount rate in an economic anal-ysis should actually reflect the risk of the project and the real ratesof return on investment obtained by private or regulated powerproducers in the same year in which the capital, O&M, and fuel costswere estimated. Since our goal is to evaluate the merits of researchand development of new technologies, we have not conducted arisk analysis and we have chosen to use the OMB suggested valueof 7%/yr for the real discount rate. To make a fair comparison toother power plant configurations, the fuel price and the discountrate were chosen to be the same in all cases. It should also be noted

that, when comparing values of LCOE between different types ofpower plants, it is important to only compare values of LCOE forprojects that produce the same type of electricity output, such aspeak-following, base-load, and the various types of intermittentoutput. Therefore, we only compare the LCOE of the Adv.IGCC–CCS and IFGC–CCS cases with other fossil energy base loadpower generation plants.

The calculation of the LCOE is significantly more challengingthan just the calculation of the system efficiency because, in addi-tion to calculating the system efficiency, one also needs to obtaincost estimates and interest rate estimates. Accurate cost estimatesare often hard to find for emerging advanced technologies thathave not been broadly commercialized, such as SOFCs, ITMs, andH2 separation membranes. While there are not many cost esti-mates for large fuel cell systems, NETL has published some esti-mates for the capital costs and replacement costs for SOFCsystems [16–18,51], for ITM [16], and for H2 separation [16,55].It should be noted that the stack costs assume mass productionSOFC stacks at the scale of roughly 500 MW installed capacityper year, and therefore the $/kW of stack capital costs used in thisand other reports are significantly lower than the $/kW of currentSOFC technology.

The other relevant figure of merit used in this paper to comparebetween power plant configurations is the internal rate of returnon investment (IRR). The IRR is the preferred figure of merit whenthe sale price is known, but the discount rate is uncertain; and it isoften the preferred figure of merit when there are two or moreproducts for sale. For example, Larson et al. [56] used an IRR anal-ysis to determine the economic viability of a biomass gasificationprocess in which there were three products for sale: steam, elec-tricity, and liquid biofuels. The average, inflation-adjusted IRR iscalculated by determining the rate of return such that the net pres-ent value (NPV) of the project equals zero.

NPV ¼ 0 ¼XN

t¼1

It

ð1þ iÞtð10Þ

where It is the net income in year t assuming no price/cost inflation,N is the total lifetime of the power plant, and i is the inflation-ad-justed rate of return. The IRR is equal to the rate of return earnedon the unrecovered balance of the investment. The value of i suchthat the NPV is equal to zero yields the geometric rate of returnon investment for this project, assuming that the yearly income isre-invested in projects with identical rates of return on investment.Hence, for the case of power plants, the IRR measures the exponen-tial nature of growth in the capability to do electrical work [57].When comparing values of IRR between different types of projects,it is important to only compare values of IRR for projects that areequally risky.

One advantage of an LCOE analysis is that the levelized capital,fuel, and labor cost can be calculated separately and summed to-gether. Another advantage is that the LCOE analysis avoids theself-referential nature of an IRR analysis; this means that the LCOEcan never yield multiple solutions. On the other hand, the advanta-ges of an IRR analysis is that it is self-referential, which is impor-tant from a public policy perspective because the IRR measuresthe estimated growth on the capital invested using today’s electric-ity prices and today’s capital, labor, and fuel costs. In this report,we calculate both the LCOE (assuming a given real discount rateof 7%) and the IRR (assuming a $50/MWh price of electricity). Insubsequent sections, we list the full details behind the LCOE andIRR calculation so that other researchers can use the cost estimates,vary some of the many inputs (fuel price, discount rate, electricityprice, etc.), and make economic comparisons with their own powerplant designs.



3.2. Economic analysis of IGCC with H2 and O2 membranes

We conducted both an IRR and a LCOE analysis of the Adv.IGCC–CCS process described earlier by making assumptions onthe capital, fuel and labor costs and then creating cash-flowtime-series for the project. The capital and labor costs for the en-trained flow gasifier, the gas turbine, the steam turbine, and cool-ing towers were averaged from the values in the IECM model [5]and those from Gerdes et al. [16,17]. The capital and labor esti-mates for the non-standard pieces of equipment were taken froma variety of sources, and will be discussed now. The estimated costof the palladium membranes was $4800/m2 of membrane surfacearea [16]. The estimated cost of the ITM ceramic membranes was$1500/m2 of membrane surface area [16]. The amount of area ofmembranes required was calculated based off the flux of H2 andO2 through laboratory scale demonstrations [27,29] of these tech-nologies at the temperatures and pressure differences chosen forthe power plant configuration. We estimated a replacement costof $8 million and $10 million every 5 yr for the palladium and cera-mic membranes, respectively, based off of the cost and amount ofthe metals and other materials required to make in the reactor. Weestimated that the cost to replace both the H2 and O2 membranes is$22.30/kW every 5 yr. We estimated that H2S/COS capture usingthe Paqell process would cost $174 million, which was based onthe cost for commercial Paqell equipment, but at 1/10th the scale[58], assuming a volume scaling factor of 0.8 for costs. The costassumptions for capital and labor are shown in Table 5 below.These costs include the cost of the equipment plus their share ofoverhead, i.e. engineering, land, construction, start-up and contin-gency. The costs associated with the coal handling are included inthe gasifier area, and the costs for 50 km of CO2 pipeline [59] areincluded in the CO2 sale/cost section, listed below.

Using the cost assumptions above as well as an average price ofbase load electricity of $50/MWh in 2007 USD and fuel price of$2/GJ for low sulfur, bituminous Appalachian coal, we were ableto determine the IRR of the project when the CO2 was sequesteredeither into an existing oil–gas well or into a saline reservoir. In theEOR case, we assumed that the CO2 could be sold at $15 per metricton of CO2 (tCO2), which is similar to the value of $12/tCO2 fromRavagnani et al. [60]. Whereas in the saline aquifer case, we as-sumed that the owners of the plant would have to pay $5/tCO2

to maintain and operate the new wells, which in the middle ofthe range of prices ($2–$7 per metric ton of CO2) estimated byEccles et al. [61]. The assumed lifetime of the plant was 25 yr, oper-ating at an 80% capacity factor, and with a 2 yr construction time.

The IRR for the EOR case is 8% per year. The IRR for the Saline caseis 3% per year.

At an inflation-adjusted discount rate of 7%/yr, the levelizedcapital costs are $25/MWh; the maintenance cost are $14/MWh;and the fuel cost is $16/MWh. This yields an overall LCOE of$58/MWh for the saline sequestration case and an overall LCOEof $47/MWh for enhanced oil recovery. The IRR and the LCOE areboth summarized in Table 6, and in addition, we include the uncer-tainty in the IRR and LCOE due to the +50%/�30% uncertainty in thecapital cost of the power plant.

3.3. Economic analysis of IGFC with catalytic gasifier and pressurizedSOFC

The system efficiency calculated earlier for this system was oneof many inputs into the economic analysis of this power plant. Todo this economic analysis, we assumed a power plant with a netelectrical output of 300 MW. In addition to the values of electricitysale price ($50/MWh) and coal price ($2/GJ), which are the same asthose listed above, some of the major inputs into the economicanalysis of this power plant are listed in Table 7. Balance of plantcosts were considered to be equal to 25% of total capital costs,and was not included in the Adv. IGCC–CCS configuration modeledpreviously because the design of the IGCC–CSS is more mature andestablished.

The non-SOFC capital costs are listed in Table 7 as capital costsdivided by the system efficiency because for many of these items,their cost decreases as the system efficiency increases. For exam-ple, both the size and cost of the gasifier island decrease as the effi-ciency of the power plant increases because a smaller gasifier isrequired to generate the same amount of net power from the plant.The SOFC cost estimates also reflect the increase in price for pres-surizing a SOFC. Since to the author’s knowledge there are no pub-lished estimates of the capital costs of a fluidized bed catalyticgasifier at commercial scale, the capital cost was conservatively as-sumed to be 50% more expensive than the GE entrained flow gas-ifier from the previous configuration model when normalized bythe same coal flow rate into the gasifier. The capital costs for thevarious pieces of equipment are listed in Table 8 when the pressureof the SOFC is 0.5 MPa.

The largest capital cost for this system was the SOFC system($915 per net kW). Of that cost, 73% is for the stacks, 16% is forthe stack enclosures, 8% is for the AC/DC converter, and 3% is fora battery system that can store 4 min of the electricity generatedat the power plant. We added the battery to the IGFC system in or-der to provide a fair comparison with PCC, NGCC, and IGCC systemswhose gas and steam turbines can provide spinning reserve for theelectrical grid. The second largest cost of the power plant was thegasifier and associated equipment ($720 per net kW).

Using the cost estimates above and assuming yearly O&M costestimates equal to 5% of the upfront capital costs, we also exam-ined how the SOFC pressure affects the IRR of this power plant con-figuration. To do so, we accounted for the change in SOFC voltageas a function of pressure as given Eq. (12). Fig. 4 shows both theexergy efficiency and the IRR of this power plant as a function ofthe SOFC pressure. The efficiency of the power plant at 2 bar is onlyroughly 40%, but at 8 bar, the efficiency is roughly 60%. Note that

Table 5Equipment capital costs including its share of overhead. O&M costs included fixed andvariable costs, assuming 80% availability. Design power output equals 808 MW. Salinecost equals $15/tCO2, and EOR sale price equals $5/tCO2, where tCO2 is metric tons ofcarbon dioxide generated.

(millions USD 2007) Initial capitalcost

Yearly O&M costs EOR (SalineSeq.)

Air separation unit(ITM)

$150 $2

Gasifier area $500 $65Particulate control $10 $2Sulfur control $170 $2WGS and H2

separation$180 $10

Power block $500 $6Water treatment $20 $5Cooling tower $100 $5CO2 sale/cost $50 �$54 ($27)Total (million $) $1680 $43 ($124)Total ($/kW) $2079 $53 ($153)

Table 6Summary of economic results for Adv. IGCC.

EOR sequestration Saline sequestration

IRR at $50/MWh 8 ± 4 (%/yr) 3 ± 3 (%/yr)LCOE at 7% real discount

rate47 ± 13 ($2007/MWh)

58 ± 13 ($2007/MWh)



the efficiency is also a strong function of the current density andthe air stoichiometric ratio, but we only analyzed the case of cur-rent density of 0.5 A cm�2 and an air stoichiometric ratio of 2.

As seen in Fig. 4, the value of IRR increases as a function of pres-sure from 0.2 MPa to 0.8 MPa (2–8 bar); however, after 0.5 MPa,there is minimal increase in the IRR. These diminished returns oc-cur because the increased efficiency with greater pressures is bal-anced by increased capital costs associated with enclosing andsealing the SOFC at higher pressure. The main reasons why wehave presented the breakdown of the capital costs (Table 8) as wellas the breakdown of the exergy destruction (Table 4) at a pressureof 0.5 MPa is that there seems to be diminishing returns above0.5 MPa and this pressure appears to be a reasonably achievablepressure for a SOFC in the near-term because 0.6 MPa is on thehigh side of pressures tested on the Rolls Royce Fuel CellSystem [64]. At a SOFC pressure of 0.5 MPa, the capital cost was$2600/kW. Due to capital cost uncertainty, a low end of values($1800/kW) represents the case in which the assumed gasifier,SOFC, and balance of plant costs have been overestimated, and ahigh end of values ($3900/kW) represents the case in which thegasifier, SOFC, and balance of plant costs have been underesti-mated. Using this range of capital costs and a price of electricityof $50/MWh, the IRR for the EOR case was 4%/yr ± 4%/yr, and theIRR for the saline aquifer case was 1%/yr ± 4%/yr. Using the rangeof capital costs and a real discount rate of 7%/yr, the LCOE for theEOR case was 52 ± 17 ($2007/MWh), and the LCOE for the salineaquifer case was 60 ± 17 ($2007/MWh), as summarized in Table 9.

4. Discussion and comparison with other economic analyses

4.1. Economic analysis of previous studies

The goal of this section is to present the capital, fuel and laborestimates of various fossil fuel power plant configurations withor without capture of carbon dioxide from previous studies, suchas Gerdes et al. [16,17], Grol and Wimer [18], and Rubin et al.[5]. These configurations include pulverized coal combustion(PCC), integrated gasification combined cycle (IGCC), natural gascombined cycle (NGCC), and integrated gasification fuel cell (IGFC).We use capital, fuel and labor estimates from these previous stud-ies, along with fuel prices and a range of possible prices of CO2

emissions in the near term, to calculate the IRR and LCOE of variouspower plant configurations. We first present the cost estimates.

Tables 10 and 11 show the first law system efficiency (%),capital costs (2007$/MWh), construction time (yr), fixed O&M($/kW/yr), availability (%), variable O&M ($/MWh), fuel cost($/MWh), and lifetime (yr) of various fossil fuel based powerplants. In addition, for the fuel cell systems listed in Table 2, thereis a reoccurring cost every 5 yr of $175/kW of generation due toSOFC stack’s replacement. The values of the maintenance/laborcosts and fuel costs can be found in Gerdes et al. [16,17], Groland Wimer [18], and Rubin et al. [5]. All cost and prices estimatesin this paper are given in 2007 USD. Using these cost and lifetimeestimates, we have generated the inflation-adjusted IRR on invest-ment for all of the major fossil-fuel based power plant configura-tions, i.e. this is the rate of return on investment assuming thatall values of price and cost inflation are exactly equal, as was donein the economic analyses presented earlier in this report.

In Fig. 5, we used the capital and labor estimates listed in Tables10 and 11 in order to calculate the IRR of various fossil fuel powerplants with and without carbon dioxide capture. We assumed afuel price of $4/GJ for natural gas and $2/GJ for coal based off of re-cent average prices coal and natural gas. As in the earlier economicanalyses, we assumed CO2 is sequestered in saline aquifers at a cost

Table 7Capital cost estimates for the Adv. IGFC configuration.

Equipment Capital cost estimation Reference

Compressor or expander $2536 (Power[kW])0.78 Extrapolatedfrom [62]

Heat exchanger $1 per cm2 of cross sectionalarea required

Extrapolatedfrom [62]

SOFC stack cost $1670 per m2 of active area Extrapolatedfrom [16,17]

SOFC enclosure $80 (p[bar])0.33 per kWgenerated in the SOFC

Extrapolatedfrom [16,17]

SOFC stack replacement $175 per kW generated in theSOFC every 5 years

Estimatedfrom [16,17]

Gasifier, solids prep/handling, and catalystregeneration

$420/(System efficiency [%])per kW of net electricitygenerated

Extrapolatedfrom [16,17]

50 km CO2 pipeline $60/(System efficiency [%]) perkW of net electricity generated

Estimatedfrom [59]

DC/AC converter $70 per kW generated in theSOFC

Estimatedfrom [16,17]

Battery $400/kWh of storage Estimatedfrom [63]

Table 8Capital cost estimate of the power plant shown in Fig. 2. SOFC pressure = 0.5 MPa,SOFC voltage = 0.7 V, SOFC current density = 0.5 A cm�2, gasifier pressure = 3 MPa,system efficiency = 58%.

Capital costs ($/kW)

Cathode air compressor 49Cathode exhaust turbine 51Syngas expander 42Syngas compressor 52CO2 compressor 17Gasifier, coal/solid prep, catalyst recovery 720CO2 capture and regeneration 120SOFC, DC/AC converter, and electrical Misc. 915CO2 pipelines 100Balance of plant 517Total 2583

Fig. 4. Exergy efficiency and internal rate of return on investment of the catalyticgasifier w/SOFC modeled in Fig. 2. The current density of the SOFC was 0.5 A cm�2.The air stoichiometric ratio was 2. Red curve is the IRR when the CO2 is used forenhanced oil recovery, and the green curve is the IRR when the CO2 is sequesteredin a saline aquifer. (For interpretation of the references to color in this figure legend,the reader is referred to the web version of this article.)

Table 9Summary of economic results for the advanced IGFC–CCS configuration at a SOFCpressure of 0.5 MPa, an air stoichiometric ratio of 2.0, and a current density of0.5 A cm�2.

EOR sequestration Saline sequestration

IRR at $50/MWh 4 ± 4 (%/yr) 1 ± 4 (%/yr)LCOE at 7% real discount rate 52 ± 17 ($2007/

MWh)60 ± 17 ($2007/MWh)



of $5/tCO2 [61] and assumed CO2 can be sold for EOR at a price of$15/tCO2 [60]. Fig. 6 shows similar information as Fig. 5, but inFig. 6, we calculate the LCOE of the various power plant configura-tions and are able to differentiate between the fuel and CO2 costs,the O&M costs and the levelized capital costs.

Of the power plants analyzed, NGCC power plants yield the high-est rate of return on investment. Though, this would only be untilthe price of CO2 emissions reaches $25/tCO2. At this price of CO2

emissions, the NGCC power plant would have an IRR of8%/yr, and this means that the following three different configura-tions would be equally viable and have an IRR of 8%: NGCC, ad-vanced IGCC–CCS–EOR, and advanced IGFC–CCS–EOR thatintegrates a catalytic coal gasifier with a pressurized SOFC. TheIRR of a conventional IGCC–100%CCS–EOR and the other IGFC con-figuration would yield an IRR near 6%/yr. Many of the configura-tions with sequestration in a saline aquifer yield a negative rateof return on investment, including the IGCC–CCS power plant con-figuration with CO2 sequestration in a saline aquifer. For the config-urations with negative values of IRR, this means that more money isspent constructing the facility than is generated in total net yearlyrevenue. Building a new PCC–CCS power plant configuration is un-likely to be economically viable compared with the alternative op-tions listed above; however, it should be noted that retrofittingexisting coal power plants may be economically viable [65], eitherfor carbon capture or for conversion into NGCC power plants.

Since a power plant requires a certain amount of useful physicalwork to be constructed, fueled, and maintained and since a powerplant also generates a certain amount of useful physical work overits lifetime, what we are attempting to express in Fig. 5 is the pre-tax, inflation-adjusted rate of return on useful physical work in-vested for various fossil-fuel power plant configurations. Whetherthese configurations with negative values of IRR, when using anelectricity sale price of $50/MWh, could achieve positive valuesof IRR at higher sale prices of electricity depends crucially on

how much the capital, fuel, and labor costs for these power plantswere to increase if the average price of electricity were to increasecompared with the price of electricity during the time period thatthe original capital, fuel and labor costs were estimated. In thisstudy, we chose a value of $50/MWh ($2007USD) because it re-flects an average base load sale price of electricity to power pro-ducers. For example, in 2007 the average price of electricity paidby industrial customers in the US was $64/MWh, respectively[66]. Since this value of $64/MWh includes transportation anddistribution costs, we have chosen to use the value of $50/MWhto reflect an average base load sale price of electricity duringthe time period that the capital, labor, and fuel prices werecalculated in this study. If the average price of electricity were toincrease in the US, such that $50/MWh were not an accurateestimate of the sale price of base load electricity, the value of IRR(in units of %/yr) could still remain the same if the percent increasein electricity prices was the same as the percent increase in capital,fuel and labor costs.

We now compare the values of IRR calculated for the Adv. IGCC–CCS and Adv. IGFC–CCS configurations analyzed in detail earlier inthis paper to the IRR of the configurations in Fig. 5. The IRR for theAdv. IGCC–CCS–EOR configuration was 8 ± 4 (%/yr) and the IRR forthe Adv. IGCC–CCS-Sal.Seq. configuration was 3 ± 3 (%/yr). Thesevalues are similar to the values of 8%/yr and 2%/yr, respectively, ob-tained using lumped cost estimates from Gerdes et al. [16,17] foran IGCC–CCS configuration with H2 and O2 separation membranesand a gas turbine operating solely on hydrogen fuel. The similarityis due to the fact that the designs were quite similar, and the costestimates for the gasifier and turbines from the IECM are similar toor are the same as the values used by Gerdes et al. [17]. The IRR forthe Adv. IGFC–CCS–EOR configuration was 4 ± 4 (%/yr) and the IRRfor the Adv. IGFC–CCS–Sal.Seq. configuration was 1 ± 4 (%/yr).These values are similar to the values of 8%/yr and 1%/yr,respectively, obtained using lumped cost estimates by Gerdes

Table 10Summary of capital, fuel and labor estimates. The capital and labor estimates are from Gerdes et al. [16,17] and Rubin et al. [5]. Fuel prices were assumed to be $2/GJ for coal and$4/GJ for natural gas. The assumed inflation-adjusted discount rate is used in the LCOE analysis presented in Figs. 6 and 7. $ = 2007 USD.

SCPCC

SC PCC–50%CCS

SC PCC–90%CCS

NGCC NGCC–90%CCS

Std.IGCC

Std. IGCC–50%CCS

Std. IGCC–90%CCS

IGCC Adv.-95%CCS

HHV efficiency 39.1% 32.9% 27.2% 49.5% 37.8% 38.2% 36.2% 32.5% 40.2%Capital cost ($/kW) 1575 2223 2870 600 1200 1813 2102 2390 2169Weighted construction time

(yr)2 2 2 1 2 2 2 2 2

Fixed O&M ($/kW/yr) 25 31 37 13 33 35 40 44 44Availability 85% 80% 80% 85% 80% 80% 80% 80% 80%Variable O&M ($/MWh) 4.9 7.1 9.4 3.0 4.0 6.5 7.3 8.1 5.3Lifetime (yr) 30 30 30 30 30 30 30 30 25Fuel ($/MWh) 17.5 20.7 25.1 27.6 36.1 17.9 18.9 21.0 17.0

Table 11Summary of capital, fuel and labor estimates for SOFC power plants in 2007 USD. The capital and labor estimates for column 1–3 are from Gerdes et al. [16,17]. For comparison,column 4 shows the estimates used Section 3.3 for modeling an advanced IGFC–CCS power plant. Fuel price was assumed to be $2/GJ for coal. The assumed inflation-adjusteddiscount rate is used in the LCOE analysis of Fig. 6. $ = 2007 USD.

Std Gasifier, 0.1 MPa SOFC Cat. Gasifier, 0.1 MPa SOFC Cat. Gasifier, 1.8 MPa SOFC Cat Gasifier, 0.5 MPa SOFC, anode recycle

HHV efficiency 42% 49% 56% 58%Capital cost ($/kW) 2135 2000 1824 2580Weight construction time (yr) 2 2 2 2Fixed O&M ($/kW/yr) 61 68 68 30Availability 80% 80% 80% 80%Variable O&M ($/MWh) 5.0 5.5 5.5 7.6Plant lifetime (yr) 20 20 20 20Stack replace time (yr) 5 5 5 5Stack replacement costs ($/kW) 175 175 175 175Fuel ($/MWh) 16.1 13.9 12.2 11.8



et al. [16,17] for an IGFC–CCS configuration with a catalytic gasifierand a pressurized SOFC. Though it should be noted that there weresome major differences between the IGFC configuration in Gerdeset al. [16,17] and the one presented here. For example, in theGerdes et al. [16,17] model there was no anode tail gas recycleand CO2 capture was accomplished via oxy-combustion of theanode tail gas. In addition, Gerdes et al. [16,17] assumed that thecatalytic gasifier costs were same as an entrained flow gasifierwhen normalized by the flow rate of coal into the gasifier, whereaswe assumed that the catalytic gasifier was 50% more expensive perflow rate of coal. This last assumption is one reason why our capitalcost estimate of the Adv. IGFC–CCS configuration ($2583/kW) is

higher than the capital cost estimate from Gerdes et al. [16,17]($1824/kW), as listed in Table 11.

In Fig. 7, we have separated out those power plants configura-tions shown in Fig. 6 that have a value of LCOE of roughly$50/MWh or less at a real discount of 7%/yr and that meet the EPA’sproposed rule of 0.45 kg (1 lb) of CO2 per kWh of electricity gener-ated [2], which was released on March 27, 2012. These configura-tions, in order of least cost to highest cost, are NGCC, Adv.IGCC–100%CCS–EOR, IGCC–50%CCS–EOR, Adv. IGFC–100%CCS–EOR (18 bar SOFC), PCC–50%CCS–EOR, NGCC–100%CCS–EOR, Std.IGCC–100%CCS–EOR, and Adv. IGFC–100%CCS–EOR (1 bar SOFC).At a price of natural gas of $4/GJ, the NGCC configuration has thelowest price of electricity. At a price of natural gas near $6/GJ, theLCOE of the NGCC power plant will be equal to the LCOE of theAdv. IGCC–100%CCS–EOR configuration. It should be noted thatwe have not analyzed any power plant retrofit configurations, andtherefore, the conclusions listed above only pertain to the econom-ics of building new power plant constructions. In the next section,we analyze the case of varying both the prices of natural gas andthe prices of CO2 emissions.

4.2. Varying the price of natural gas and CO2 emissions

The results presented in Fig. 5 suggest that the IRR of an ad-vanced IGCC–CCS power plant is similar to the IRR of an advancedIGFC–CCS if the catalytic gasifier costs are the similar to the cost ofan entrained flow gasifier for similar input of coal and if SOFC tech-nology can achieve mass production cost targets. However, the val-ues of IRR calculated for these advanced power plantconfigurations were well below the IRR of a NGCC power plantwith today’s fuel prices and no CO2 tax. We now address the fol-lowing question: at what price of natural gas and at what priceof CO2 emissions would an advanced IGCC–CCS–EOR orIGFC–CCS–EOR power plant configuration be economically viable?We therefore conducted an LCOE analysis for the various powerplant configurations as a function of the price of natural gas andcarbon dioxide emissions, holding all other variables constant.

Figs. 8 and 9 show which fossil fuel power plant configurationhas the lowest LCOE as a function of the cost of natural gas (NG)and the cost of emitting CO2 into the atmosphere while holdingthe cost of coal at $2/GJ. In Fig. 8, we assume that the captured

Fig. 5. Internal rate of returns on investment (IRR) of building new fossil fuel, baseload power plants. Data in blue represents the rate of return if there is no tax forCO2 emissions. Data in orange represents the rate of return if there is tax foremissions of $10/tCO2. Data in red represents the rate of return if there is tax foremissions of $20/tCO2. Data in green represents the rate of return if >90% of the CO2

is sequestered in saline aquifers at a cost $5/tCO2. Data in gray represents the rate ofreturn if >90% of the CO2 is sold for enhanced oil recovery at a price of $15/tCO2. Thefuel price for coal was assumed to be $2/GJth and the natural gas fuel price was$4/GJth. The sale price of electricity was assumed to be $50/MWh, i.e. $14/GJel.$ = 2007 USD. (For interpretation of the references to color in this figure legend, thereader is referred to the web version of this article.)

Fig. 6. Levelized cost of electricity (LCOE) in 2007 USD/MWh of building new fossilfuel, base load power plants. The cost is broken down into levelized capital costs,the fixed plus variable O&M, and the sum of the cost for fuel plus CO2 emissions orsales. ‘Sal. Seq.’ means sequestration in saline aquifers at a cost $5/tCO2. ‘EOR’means enhanced oil recovery at a positive sale price of $15/tCO2. The fuel price forcoal was assumed to be $2/GJth and the natural gas fuel price was $4/GJth. Theassumed inflation-adjusted discount rate was 7%/yr. Capital and O&M costestimates are listed in Tables 1 and 2.

Fig. 7. Levelized cost of electricity (LCOE) in 2007 USD/MWh of building new fossilfuel, base load power plants. The configurations above meet proposed EPAregulations of 0.45 kg (1 lb) of CO2 per kWh of electricity generated. The cost isbroken down into levelized capital costs, the fixed plus variable O&M, and the sumof the cost for fuel plus CO2 EOR sales at $15/tCO2. The fuel price for coal wasassumed to be $2/GJth and the natural gas fuel price was $4/GJth. The assumedinflation-adjusted discount rate was 7%/yr. Capital and O&M cost estimates arelisted in Tables 1 and 2.



CO2 can be used for EOR, whereas in Fig. 9, we assume that the cap-tured CO2 must be sequestered in a saline aquifer. For the EOR case,we found that if the price of natural gas goes above the line be-tween the points ($10/tCO2, $5.0/GJ) and ($50/tCO2, $2.5/GJ), thenthe Adv IGCC & IGFC–CCS–EOR configurations have the lowest lev-elized cost of electricity. Note also that there is line at which aNGCC–CCS power plant has the lowest value of LCOE and there isalso a vertical line around $10/tCO2 at which a PCC has the lowestvalue of LCOE. There is also a horizontal line just below $2.5/GJ,which shows when the Adv. IGCC and IGFC–CCS configurationsand NGCC–CCS configurations have the same LCOE. For salinesequestration (ca. Fig. 9), the area in the graph in which NGCChas the lowest value of LCOE increases substantially. The resultsin Fig. 8 are fairly similar to the results presented in Fig. 3 of Fisch-beck et al. [12], who present a case in which captured carbon diox-ide is assumed to obtain a sale price of $25/tCO2. However, onemajor difference is that our ‘Coal with CCS’ case was an Adv.IGCC–CCS configuration rather than a PC–CCS configuration.

The economic analyses conducted here suggest that there mightbe scenarios in which Adv IGCC and IGFC–CCS power plant config-urations are economically preferable; however, this requires eitheran increase in the price of natural gas or an increase in the price ofCO2 emissions. It should also be noted that this analysis held theprice of capital, labor, and coal constant while varying the price of

natural gas and the price of CO2 emission. The price of capital, labor,and coal is unlikely to remain constant with changing price ofnatural gas and CO2 emission, so the conclusion we drawn fromthese figures should not be used as predictions for future outcomes.Instead, they should be used to determine which power plantconfigurations deserve further research and development. It shouldbe noted that the advanced IGCC and IGFC configuration studiedhere still require significant levels of research and developmentbefore they are commercially-viable for large scale power plants.Specifically, this means scaling up H2 and O2 membrane technology,further testing operation of gas turbines on hydrogen, scaling upthe size of the catalytic gasifier, proving the catalyst regenerationprocess, and scaling up the size and pressure of SOFCs.

5. Conclusions

We conducted exergy and economic analyses for two advancedcoal-based power plants with CO2 capture and sequestration.When conducting our capital cost analysis for the advancedIGCC–CCS–EOR design, we calculated a system efficiency of 43%and a value of IRR of 8 ± 4%/yr (at $50/MWh), which was similarto the value estimated in previous studies [16,17]. When conduct-ing our capital cost analysis for the advanced IGFC–CCS–EOR de-sign, we calculated a system efficiency of 58% and a value of IRRof 4 ± 4%/yr (at $50/MWh), which was also close to the values esti-mated by previous researchers. We used capital and labor cost esti-mates from previous researchers in order to compare the IRR ofthese two advanced power plant designs with conventional andother advanced power plant designs. Using cost estimates fromother studies and assuming recent fuel and electricity prices, a nat-ural gas combined cycle (NGCC) power plant yielded the highestvalue of rate of return on investment (22%/yr). However, our re-sults suggest that, in the case of a CO2 tax of $25/tCO2, thenthree different configurations are equally viable economically(IRR = 8%/yr at $50/MWh): a NGCC power plant without capture,an advanced IGCC–CCS–EOR power plant with H2 and O2

membranes, and an advanced IGFC–CCS–EOR power plant thatintegrates a catalytic coal gasifier with a pressurized SOFC.

This research suggests that there may be advanced coal-basedpower plants that can achieve reasonable values of IRR (8%/yr) attoday’s typical prices for base load electricity generation of $50/MWh; and therefore, research into advanced H2 and O2 separationmembranes as well as pressurized SOFCs are of crucial importanceto the development of low cost base load electricity if the price ofnatural gas in the future goes above $5/GJ and the price of CO2

emissions goes above $20/tCO2. The calculations in this report sug-gest that there may be scenarios in which advanced IGCC and IGFCconfiguration are economically viable, meriting further researchand development into these technologies.

Acknowledgments

We thank Dhruv Bhatnagar and Srikant Subramaniam for theircontributions to this work. We thank Drs. Jay Apt, David Berry, PhilDiPietro, Paul Fischbeck, Kristin Gerdes, Dale Keairns, David Rode,Edward Rubin, Dushyant Shekhawat, Wayne Surdoval, and JanThijssen for their discussions on this topic. We thank the financialsupport of the National Energy Technology Laboratory as part ofthe US Department of Energy.

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Fig. 8. The lowest cost of electricity between PCC, NGCC, NGCC–CCS–EOR, Adv.IGFC–CCS–EOR, and Adv. IGCC–CCS–EOR, as a function of the price of natural gasand CO2 emissions, assuming an EOR sale price of $15/ton of CO2, using costestimates from Gerdes et al. [16,17]. The price of coal was held constant at $2/GJ in2007USD.

Fig. 9. The lowest cost of electricity between PCC, NGCC, NGCC–CCS-SalSeq, Adv.IGFC–CCS-SalSeq, and Adv. IGCC–CCS-SalSeq, as a function of the price of naturalgas and CO2 emissions, assuming a Saline Sequestration cost of $5/ton of CO2, usingcost estimates from Gerdes et al. [16,17]. The price of coal was held constant at$2/GJ in 2007USD.



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